Hydrocracking nitrogen containing feedstocks

ABSTRACT

HYFROCARBONS CONTAINING SUBSTANTIAL AMOUNTS OF ORGANO-NITROGEN COMPOUNDS ARE CONVERTED IN THE PRESENCE OF HYDROGEN OVER A CATALYST COMPRISING ALUMINA, ALUMINOSILICATES, OR COMBINATIONS THEREOF WITH SILICA, ZIRONIA, MAGNESIS AND THE LIKE WITH OR WITHOUT ACTIVE METAL COMPONENTS IN THE PRESENCE OF CONTROLLED MINOR AMOUNTS OF ADDED WATER.

Patented Oct. 17, 1972 3,699,036 HYDROCRACKING NITROGEN CONTAININGFEEDSTOCKS Robert H. Hass, Fullerton, and Cloyd P. Reeg, Orange, Calif,'assignors to Union Oil Company of California, Los Angeles, Calif. NoDrawing. Filed Aug. 21, 1970, Ser. No. 66,062 Int. Cl. Cg 13/02, 13/04,13/10 US. Cl. 208-111 14 Claims ABSTRACT OF THE DISCLOSURE Hydrocarbonscontaining substantial amounts of organo-nitrogen compounds areconverted in the presence of hydrogen over a catalyst comprisingalumina, aluminosilicates, or combinations thereof with silica,zirconia, magnesia and the like with or without active metal componentsin the presence of controlled minor amounts of added water.

BACKGROUND Recent years have witnessed a phenomenal growth in thedevelopment and application of catalytic hydrocarbon conversionprocesses including hydrocracking, hydrofining, hydrogenation andcombinations of these. By far the greatest part of these eiforts havebeen directed toward methods for hydrocracking gas oils to productsboiling in the gasoline range. The best catalysts developed to date forthis purpose are those comprising a highly active cracking base, e.g.,of the crystalline zeolite type, combined with a highly activehydrogenation component such as palladium or platinum. These catalystsare'highly efiicient for converting middle distillate oils boiling inthe 400-750 F. range to gasoline. Due to the extensive present andprospective use of middle distillate stocks as feed to thesegasoline-producing hydrocracking units, and to other economic,geographic and seasonal factors, a need is being felt in the industry toprovide additional middle distillate stocks to meet the demand for otherproducts such as turbine and diesel fuels. An obviously desirable sourceof such additional middle distillate stocks comprises the heavydistillates boiling above about 700 F., which have heretofore beendiverted largely to fuel oils because of the lack of an economicalmethod for converting them to lower boiling products.

In the initial investigation of hydrocracking techniques for convertinghydrocarbon feeds boiling above 700 F. it became apparent that severehydrocracking to produce gasoline in a single pass conversion step wasimpractical, firstly because inordinate amounts of butanes and dry gaswere produced, and secondly because the gasoline product was always ofundesirably low octane quality. It thus became apparent that forpurposes of gasoline production, especially on a once through basis,hydrocracking techniques were of primary value in connection with middledistillate feedstocks, and that the hydrocracking of higher boilingstocks would be economically desirable only if conversion to gasolinecould be minimized and the production of middle distillates maximized. Aprimary objective in the hydrocracking of these heavy feeds was thus therealization of maximum middle distillate/ gasoline ratios in theresulting product.

It was then found that the catalysts most useful for converting middledistillates to gasoline were least useful for converting heavyfeedstocks selectively to middle distillates in that large amounts ofdry gases, butanes and gasoline were produced by those catalysts. Agreat many conventional hydrocracking catalysts were tested in anattempt to find one which would convert such feedstocks more selectivelyto middle distillate products. In all cases it was found that thedesired selectivity could be maintained only by operating at very lowoverall conversion per pass, entailing prohibitively high recycle oilrates. However, upon developing and testing certain catalysts notconventionally regarded as hydrocracking catalysts, substantially morepromising results were obtained. Specifically, it was found that byusing a type of catalyst commonly employed for catalytic hydrofiningcomposed of nickel and molybdenum oxides and/or sulfides supported onactivated alumina, the selectivity of conversion to middle distillateproducts was excellent even at relatively high conversions, e.g., 40-60volume-percent per pass. Moreover, for feedstocks containing organicnitrogen, it was found that these catalysts were not only moreselective, but sometimes more active (on the basis of a lowertemperature required for a given total conversion) than even the mostactive hydrocracking catalysts based on zeolite cracking bases.

Although those catalysts are particularly attractive for convertingrelatively high boiling nitrogenous feeds to middle distillates and highquality midbarrel fuels, they also have considerable utility in theconversion of similar feedstocks to high octane gasoline boiling rangehydrocarbons. However, as is always the case with chemical conversionsystems, there remains considerable room for improvement in numerouslimiting process parameters such as catalyst activity, productdistribution, catalyst life and the like. For example, it is known thatcatalyst life can be dramatically increased by lowering reactiontemperature. However, such procedures decidedly limit the obtainableconversion'per pass requiring greater capital and operating expense dueto the requirement of larger reactor sizes and recycle systems.Therefore, a reasonable balance must be drawn between catalystdeactivation rate, reaction temperatures and tolerable conversion levelsthat enable efiicient use of process equipment.

Yet another consideration which limits the effectiveness of thesecatalyst systems for hydrocarbon conversion, particularly hydrocrackingand hydrogenation, is the deleterious effect exhibited on catalystactivity by organonitrogen compounds in the feed. Although some of thesecatalysts are particularly elfective for hydrofining, e. g.,denitrogenation and desulfurization, we have observed that theirhydrocracking and hydrogenation activity is markedly limited by thepresence of even minor amounts of organonitrogen compounds when theconcentration of those materials in the feed substantially exceeds 1p.p.m. nitrogen. The most direct approach to the solution of thisproblem would appear to be the use of a hydrofining zone upstream of thehydrocracking system or the use of a larger hydrocracking catalyst bed.In these systems, and in other apparent alternative schemes, sufficientprecontacting of the nitrogen containing feed with these or otherhydrofining catalysts under hydrofining conditions can be effected toreduce the nitrogen content to tolerable levels, e.g., less than about 5p.p.m. Although such procedures undoubtedly provide an availablealternative, they are subject to certain obvious disadvantages. In orderto reduce the nitrogen content of the feed to a level sufficient toprevent inhibition of hydrogenation and hydrocracking, provision must bemade for rather severe hydrofining treatment upstream of the contactingzone. Pretreatments of this nature are known to involve considerexpense.

Although organonitrogen compounds are known to exhibit severaldeleterious eifects on the preferred midbarrel hydrocracking andhydrogenation catalysts such as those comprising Group VI and Group VIIImetal components on a primarily alumina support, their influence on theactivity, selectivity and elfective catalyst life of the zeolite basedcatalyst particularly those containing Group VIII noble metals such asplatinum and palladium, is even more dramatic. For example it is knownthat with some of the zeolite catalysts primarily employed for gasolineproduction that the presence of substantially more than parts permillion of organonitrogen compounds in the feed determined as elementalnitrogen almost completely inhibits the hydrocracking and hydrogenationactivity of those compositions, at least in the upstream sections of thecontacting zone, until sufiicient contacting has been effected to reducethe organonitrogen content to tolerable levels. As a general rule, thezeolite base catalysts are not particularly effective denitrogenationcatalysts and many are permanently inhibited by exposure to substantialamounts of organonitrogen compounds. As a result, it is a matter ofalmost necessity in the use of such composition to provide for thealmost complete 'den-itrogen'ation of a selected feedstock prior tohydrogenation or hydrocracking over these catalysts. It is usuallynecessary to reduce the organonitrogen content of such feeds to a pointbelow about 5'p.p.m. and inmany cases, nitrogen levels substantially inexcess of 1 p.p.m. cannot be tolerated. Obviously, the expense andprocess complexities involved in effecting these degrees ofvdenitrogenation are considerable. However, to date they have beenconsidered acceptable in view of the lack of available alternatives andthe relative degree of comparative success realized by operating in thatmanner.

We have now discovered a procedure whereby several aspects of suchhydrocarbon conversion systems can be markedly improved from thestandpoint of performance and economy. This procedure enables theattainment of economically feasible conversion levels and productquality at less severe reactor conditions, e.g., lower reactiontemperatures, without the necessity of severe prehydrofining. We havediscovered that the addition of controlled amounts of water or waterprecursors to the hydrocarbon feed counteracts the effect of substantialnitrogen levels and enables the catalysts to effect acceptableconversion levels at reduced temperatures. Several of the advantages ofthese procedures are readily apparent. Firstly, longer catalyst life canbe realized due to the lower reaction temperatures required. Secondly,operating and capital expense involved in upstream hydrofining zones canbe reduced due to the tolerance of the described conversion systems tohigher concentrations of nitrogen compounds. In addition, the activityof these systems for several forms of hydrocarbon conversion issubstantially improved, thereby enabling higher conversions, loweroperating temperatures, greater control of product distribution, and thelike. Particularly dramatic efi'ects are evidenced in hydrocrackingactivity and hydrogenation, particularly the hydrogenation of aromaticconstituents. Consequently, these systems can be operated underconditions which provide markedly superior results in the production ofthose gasolines and midbarrel fuels and even enable the production ofsuch products from relatively highly aromatic refractory feedstocksunder recycle conditions while minimizing the degree of refractorybuildup in the recycle stream.

It is therefore one object of this invention to provide an improvedhydrocarbon conversion process. It is another object of this inventionto provide an improved process for producing middle distillate fuelsand/ or gasoline. It is yet another object of this invention to providea method for increasing the activity of hydrocracking systems. Yetanother object of vthis invention is the provision of a method forincreasing the tolerance of hydrocracking and hydrogenation systems toorganic nitrogen compounds. Still another object of this invention isthe provision of improved hydrogenation, hydrocracking and hydrofiningmethods. Another object of this invention is the provision of a methodwhich enables the reduction of hydrocracking temperatures necessary toobtain prescribed conversion levels.

In accordance with one embodiment of this invention, hydrocarbonscontaining at least about 5 p.p.m. nitrogen as organonitrogen compoundsare contacted in the presence of hydrogen under hydrocarbon conversionconditions with a catalyst comprising at least one of the metals, oxidesand sulfides of Groups II, IV-B, VI-B, VII-B and VIII combined with aforaminous oxide support in the presence of a promoting amount of addedwater or water precursor.

In accordance with another embodiment of this invention, hydrocarbonscontaining at least about 5 p.p.m. nitrogen as organonitrogen compoundsare contacted in the presence of hydrogen at elevated temperatures andpressures sufficient to promote the reactions of said hydrocarbons withthe added hydrogen in the presence of a catalyst comprising at least onehydrogenation active metal, oxide or sulfide of Groups VI-B, VII-B andVIII combined with alumina, silica, magnesia, zirconia or clayscomprising primarily alumina and/or silica and combinations thereof suchas aluminosilicates in the presence of a controlled added amount ofwater equivalent to at least about 0.2 weight-percent and based onhydrocarbons.

During the investigation of this phenomena water was supplied to ahydrocracking zone at several different sets of operating conditions byseveral different procedures. In one instance water was admixed with ahydrocarbon phase by humidifying the hydrogen stream to the reactor. Inanother instance, the same results were effected by adding a Waterprecursor, e.g., n-butanol to the feed. Alcohols of this type arebelieved to be dehydroxylated to produce water under hydrocrackingconditions. In every case, the addition of Water by these proceduresimproved the activity of the system. However, an even more surprisingresult was observedwhen the water injection was discontinued. Instead ofthe catalyst deactivating gradually to its original level as expectedthe hydrocracking performance exhibited a transient behavior in areproducible cycle during which the activity improved still furtherreaching a maximum and then slowly declined to the activity exhibited bythe system prior to water addition. This transient phenomenon wasreversible and repeatable. However, despite all efforts, includingperiodic water injection, no means could be found whereby the maximumcatalyst activity observed could be sustained on a continuous basis.

Although the exact mechanism by which this phenomenon is effected is notunderstood with certainty, it may be that the presence of watermitigatesthe inhibiting effect of organic nitrogen compounds by either preventingaccess of those poisons to the catalyst surface or by increasing thetolerance of the catalyst to those impurities. These suppositions aresupported somewhat by observations made during operations on hydrocarbonfeedstoeks having very low nitrogen concentrations, i.e., below about 1p.p.m. organic nitrogen. In those systems, the catalystactivity was seento decline upon water addition as would be expected since it isgenerally believed that water itself is a catalyst inhibitor.Nevertheless, in view of these observations, it seems reasonable toconclude that the phenomena herein described is at least in part afunction of the combined effects of water and nitrogen.

The activity exhibited by these catalyst systems at a given set ofprocess conditions is greater in the absence of either nitrogen orwater. Both of these materials are generally considered catalystinhibitors in otherwise clean systems. However, we have observed thatconsiderable advantage can be realized by the addition of water tohydrocracking zones operating on hydrocarbon mixtures containingsubstantial amounts of organonitrogen compounds, i.e., 5 p.p.m. orgreater.

Those skilled in the art readily recognize that the prehydrofiningseverity required to reduce the nitrogen content of hydrocarbon mixturesto levels below 5 p.p.m. is not easily attained and involvesconsiderable expense. The method of this invention offers an alternativeto such severe hydrofining procedures. It is possible to drasticallyreduce the severity and expense of prehydrofining treatments andcompensate for the residual nitrogen in the hydrocarbon feed byconducting the conversion in the presence of controlled amounts of wateras described. In fact, the advantages of this invention can be realizedin the absence of any hydrofining whatever, provided of course that thenature of the feedstock is not such that it completely destroys theactivity of the catalyst.

A wide variey of hydrocarbon feeds can be employed within the concept ofthis invention. Similarly, the conditions under which the feedhydrocarbons are reacted and the compositions of the catalysts employedin any specific case can be varied to promote selected reactionmechanisms such as hydrogenation, hydrofining, hydrocracking, and thelike. The advantages of this invention can be realized in all of thesesituations when the feed employed contains a substantial amount oforganonitrogen compounds. As a general rule, the hydrocarbon feedemployed in these systems will boil above about 200 F. and will includelight, medium and heavy naphthas, gas oils, coker distillates, vacuumgas oils, and the like. These systems may be designed for the simplehydrogenation of aromatic feedstock of essentially any molecular weight.However, particular advantage is realized in the hydrogenation ofaromatic feedstocks particularly those boiling above about 400 F. andcontaining in excess of about and usually about to about 200 p.p.m.nitrogen as organonitrogen compounds which we have found to markedlyinhibit the aromatics hydrogenation activity of the catalystsencompassed within the scope of this invention. However, the presentlymost preferred application of this phenomena is in the hydrocracking ofraw or partially refined hydrocarbon feeds to produce gasolines andmidbarrel fuels. We have observed that the addition of the describedcontrolled water levels to hydrocracking systems dramatically improvesthe activity thereof and enables much greater control of productdistribution than is otherwise possible. The feedstocks employed in suchhydrocracking systems boil above about 400 F., usually within the rangeof about 400 to about 1200" F. Depending upon the degree of prerefining,the organonitrogen content of these feeds can vary substantially.Organonitrogen compounds are usually present in amounts equivalent toabout 10- p.p.m. to about 1 weight-percent nitrogen, generally about 10to about 5000 p.p.m. nitrogen. However, we have observed that theadvantages of operating in the described manner are more apparent whenthe nitrogen content of the feed is within the range of about 10 toabout 600 ppm. Consequently, it is presently preferred that feedstockshaving nitrogen levels substantially in excess of 1000 p.p.m. bepartially hydrofined prior to contacting in the hydrocracking systemsoperated in accordance with this invention.

Although the sulfur content of selected feedstock is not presentlyconsidered to be a determinative variable with regard to the resultsobserved upon water addition, it is nevertheless a considerationrelative to overall catalyst activity. The deactivating effect ofexcessive sulfur concentrations can be substantially mitigated by theuse of sulfactive catalysts, which are generally well known in the art.Similarly, the presence of some sulfur compounds in the feed is oftenpreferred in hydrocracking systems in that these materials tend tosulfide the catalyst maintaining it in the generally more activesulfided state. As a general rule, the raw hydrocarbon feeds may containorganic sulfur concentrations ranging up to about 3 weight-percentsulfur based on the total feed. It is often preferable to substantiallyreduce this sulfur content to a level below about 100 ppm. bypre'hydrofining. Such desulfurization can be conveneiently effected bycontacting at conventional hydrofining conditions in the presence ofcatalysts similar to those herein described. The hydrogen sulfideresulting from the desulfurization of the organic sulfur compounds canbe separated prior to introduction of the feed to the hydrocracking zoneor can be passed directly over the hydrocracking or hydrogenationcatalyst without substantial detriment.

A wide variety of catalysts can be employed within the concept of thisinvention. Essentially any catalyst comprising an active form of a metalor metal compound of Groups H, VI, VII and VIII of the Periodic Chartsupported on a foraminous refractory oxide can be employed in thesesystems. Since catalysts of this nature are generally well known in theart they need not be described in detail herein. However, briefreference will be made to several of the presently most popularcompositions within this class.

The foraminous refractory oxides above referred to can be selected fromalumina, silica, magnesia, zirconia and alumina and silica containingclays such as bentonite and attapulgus clay and combinations thereofsuch as silica, alumina, silica zirconia and the like. One presentlypreferred combination of these materials is silica stabilized aluminacontaining about 2 to about 30 weightpercent silica. This material initself is conventional in the art and can be prepared by severalprocedures such as cogelation or sequential gelation or physicaladmixture of partially or completely precipitated gels of alumina andsilica. It has been observed that the presence of minor amounts ofsilica in alumina matrices greatly improves the physical properties ofthe resultant combination. The natural and synthetic crystalline andamorphous aluminosilicate Zeolites are also widely employed for thepreparation of hydrocarbon conversion catalysts and are applicable tothe systems herein described.

Exemplary of these compositions are the aluminum containing ionexchangeable clays, particularly acid extracted clays such as bentonite,montmorillonite, beidellite, halloysite, endelite, kalolinite and thelike. The acid extraction of these materials is generally well known inthe art and is designed to effect the removal of metals found in thealuminosilicates in their naturally occurring state. Such acidextraction can be effected by contacting the clay with a mildly acidicsolution or a strong mineral acid such as nitric, sulfuric,hydrochloric, or phosphoric acids and the like, having a pH of about 4or above.

A second generally recognized class of amorphous aluminosilicates havingthe above-described characteristics are the partially degradedcrystalline aluminosilicates which have been subjected to acidic and/orthermal en vironments sufficient to at least partially destroy thecharacteristics crystalline structure of those materials. As a generalrule, there is little economic interest in obtaining starting materials,i.e., aluminosilicates, by this latter procedure in that the crystallinealuminosilicates are usually far more expensive than alternative formsof aluminosilicates suitable for use herein. Nevertheless, suchmaterials are suitable for application within the concept of thisinvention.

Another class of aluminosilicates having the desired characteristics arethe silica-alumina cogels. These materials are usually prepared byeither coprecipitating silica and alumina from an aqueous solution atsoluble salts of silicon and aluminum or by grafting one of theseconstituents, i.e., either the silica or alumina, onto a previouslyprecipitated alumina or silica gel by acidifying a solution of the watersoluble salt of the second component. For example, alumina can beprecipitated in the presence of a hydrous silica gel by acidifying asolution of sodium aluminate with a mineral acid such as sulfuric,nitric, and the like. The relative amounts of silica and alumina in thecombinations can vary considerably although it is presently preferredthat the silica concentration be equivalent to about 10 to about 40weight-percent of the total silica-alumina combination on a dry weightbasis. The cogels resulting from such coprecipitation or sequentialprecipitation of separate constituents can then be dried and ionexchanged to replace the undesirable cations with ammonium, hydrogenand/or Group VIII metal-containing cations as previously described.

The presently preferred aluminosilicates are the crystalline specieshaving Slo /A1 ratios of at least about 2. This class includes bothsynthetic and naturally occurring zeolites. Illustrative of thesynthetic zeolites are zeloite X, U.S. 2,882,244; zeolite Y, U.S.3,130,007; zeolite A, U.S. 2,882,243; zeolite L, Belg. 575,117; zeoliteD., Can. 611,981; zeolite R U.S. 3,030,181; zeolite S, U.S. 3,054,657;zeolite T, U.S. 2,950,952; zeolite Z, Can. 614,995; zeolite E, Can.636,931; zeolite F, U.S. 2,995,358; zeolite O, U.S. 3,140,252 zeolite B,U.S. 3,008,- 803; zeolite Q, U.S. 2,991,151; zeolite M, U.S. 2,995,-423; zeolite H, U.S. 3,010,789; zeolite J, U.S. 3,011,869; zeolite W,U.S. 3,012,853; zeolite KG, U.S. 3,056,654. Illustrative of thenaturally occurring crystalline aluminosilicates which can be suitablytreated by the methods herein described are levynite, dachiardite,erionite, faujasite, analcite, paulingite, noselite, ferrierite,haulandite, scolecite, stilbite, clinoptilolite, harmotome, phillipsite,brewsterite, flakite, datolite, chabazite, gmelinite, cancrinite,leucite, lazurite, scolacite, mesolite, ptilolite, mordenite, nepheline,natrolite, and sodalite. The natural and synthetic faujasite-typecrystalline aluminosilicate zeolites, e.g., zeolites X and Y,arepreseutly particularly preferred.

When these catalysts are prepared in whole or in. part from zeoliticaluminosilicates, it is often desirable to substantially reduce thealkali metal content of the aluminasilicate during some stage of thecatalyst preparation and convert the zeolite into an acidic form.Consequently, the alkali metals which are present both in natural andsynthetic aluminosilicates in their original state are preferablyremoved during some stage of the catalyst preparation by ion exchange.Any one of numerous procedures can be employed for this purpose.Probably the most convenient procedures include ion exchange withaqueous solutions of ammonium or hydrogen or soluble salts of metalcations of Groups II, VII and VIII of the Periodic Chart. As a generalrule, the alkali metal content of the aluminosilicates should be reducedto less than 5 weightpercent preferably less than about 2 weight-percentdetermined as a corresponding oxide. This objective can be realized bycontinuing the exchange with the cations above mentioned until thedesired reduction in alkali metal content has been achieved. Whenexchange with the noted metal cations is employed, the resultantconcentration of the respective metals in the final composition willusually be within the range of about 0.2 to about 5 weight-percentdetermined as the elemental metal. These metals can also be added to thealuminosilicate or other foraminous oxides by impregnating the supportwith an aqueous solution of a water soluble thermally deco-mp0sable saltof the desired metal. The Group VI-B metals can also be added by thisprocedure. Illustrative of suitable water soluble thermally decomposablecompounds are the nitmtes, sulfates, carbonates and halides of calcium,magnesium, molybdenum, tungsten, manganese, rhenium, nickel, cobalt,platinum, palladium and the like.

Catalysts presently particularly preferred for the production ofmidbarrel fuels comprise composites of nickel and/ or cobalt metals,oxides and sulfides and one or more of molybdenum and tungsten metals,oxides and sulfides combined with an aluminum support.

The aluminas presently preferred are high surface area amorphousmaterials having relatively low cracking activities corresponding toCat-A indices below 25. However, higher activities can be tolerated whenhigher relative gasoline production is desired. It is also preferredthat the alumina support be stabilized by combination with a minoramount of amorphous silica gel. Silica can be present in concentrationsof up to about 40 weightpercent based on the combined weight of silicaand alumina. Presently preferred compositions have silica concentrationswithin a range of about 3 to about 25 weight-percent.

The presently preferred midbarrel catalyst compositions contain about0.5 to about 10 Weight-percent nickel and/or cobalt metals, oxides orsulfides determined as the corresponding oxides, and about 4 to about 30weightpercent of molybdenum and/or tungsten metals, oxides or sulfidesbased on the corresponding oxides. Combinations of nickel and molybdenumsulfides are particularly preferred. Such compositions are normallyprepared by impregnating the selected alumina carrier with an aqueoussolution or solutions of soluble salts of the respective metals followedby draining, drying and calcining in air at temperatures of 800 to about1200 F. The calcined catalysts are preferably presulfided prior tocontact with a feedstoc k as by reaction with a gaseous mixture ofhydrogen and hydrogen sulfide, carbon disulfide, elemental sulfur andthe like.

The preferred carrier for these midbarrel fuels catalysts is activatedalumina gel containing a minor proportion of coprecipitated silica gel.The silica content should not exceed about 40 percent by weight. Highersilica contents tend to decrease cracking selectivity to midbarrelfuels. The preferred silica content is between about 3 and 25 percent byWeight. The preferred supports comprise at least about 60 weight-percentalumina, the remainder being silica, magnesia, aluminosilicates, and thelike. Prior to impregnation, the carrier is preferably formed intopellets of about 4 to inch diameter (by extrusion or die-compression).Alternatively, the catalyst may be employed in a powder form.

Particularly preferred catalyst compositions are prepared byimpregnation with the stabilized impregnating solutions discussed inU.S. Pats. 3,232,887 and 3,287,280. By these procedures the selectedsupport is contacted with a solution of the desired metal saltscontaining stabilizing amounts of phosphate ions introduced by theaddition of an acid of phosphorus, e.g., phosphoric acid, phosphorousacid and the like. Phosphorus concentrations in the final catalyst areat least about 0.2 weightpercent, usually within a range of about 0.5 toabout 8 weight-percent, preferably about 2 to about 5 weightpercentbased on the free metal.

Hydrocracking conditions generally considered to be most desirableinclude reaction temperatures Within a range of 500 to about 900 F.,preferably 550 to 850 F., reactor pressures of at least about 500p.s.i.g., preferably 1000 to about 5000 p.s.i.g., and liquid hourlyspace velocities (LHSV) of about 0.1 to about 10, preferably 0.3 toabout 5. The free hydrogen content in the reaction zone should beequivalent to at least about 5 00, generally about 500 to about 20,000and preferably about 1000 to about 15,000 standard cubic feet per barrelof reactor charge.

Hydrogenation and hydrofining, e.g. denitrogenation and desulfiurizationcan be effected at these same conditions and in fact, are etfected underhydrocracking conditions when there are constituents in the feed subjectto hydrogenation and/hydrofining. However, it is sometimes desirable toeffect denitrogenation, desulfurization and/ or hydrogenation in theabsence of substantial degrees of hydrocracking. In those instances, itis desirable to contact the feed under somewhat milder conditions ofreduced temperature and pressure, e.g., pressures-of about to about 1200p.s.i.g. Catalyst compositions also determine the relative degrees ofhydrocracking and hydrogenation and can be selected to favor one overthe other as desired.

The equivalent water concentration in the reaction zone should besnfiicient to promote the activity of the catalyst as evidenced byhigher conversion levels at otherwise identical conditions. We havefound that the most dramatic improvement in catalyst activity iseffected at equivalent water concentrations in excess of about 0.1weight-percent, generally about 0.2 to about 5 weight-percent andpreferably about 0.2 to about 2 weight-percent. Depending upon thecharacteristics of the particular systems such as reaction conditionsand the selected catalyst and the organo nitrogen content of theselected feedstock, the preferred water level in the reaction zone may-vary considerably from about 0.1 to about weight-percent. A constantimprovement in activity can be realized by maintaining a predeterminedcontrolled water concentration in the reaction zone. We have observedthat the rate of the increase in catalyst activity is related to themaximum equivalent water concentration in the reaction zone, certainreaction parameters such as liquid hourly space velocity and mixingcharacteristics and of course the duration of water injection into thereaction zone. Therefore, it is possible when employing the cyclic waterinjection procedure previously described, to extend the cycle lengthduring which these beneficial effects are realized by operating at thehigher water concentrations and/or lower liquid hourly space velocities.

An essential aspect of this procedure is that the prescribed controlledwater concentration be provided in the reaction zone. This can beaccomplished by the addition of either free water or a water precursorto the hydrocarbon feed, to one or more of the streams entering thereactor, or directly to the reactor. Exemplary of readily availablewater precursors are the organic alcohols. Presently preferred due totheir availability and cost are those having from about 2 to about 12carbon atoms and one to about five hydroxyl groups per molecule.

Although it is presently preferred that the water concentrations bemaintained only in the hydrocracking zone, it is within the scope ofthis invention to provide for the addition of water to process systemsupstream of the hydrocracking zone. The most apparent alternative inthis regard is the addition of water or water precursor to thehydrocarbon feed entering a hydrofining stage immediately preceding thehydrocracking zone. This mode of operation is particularly convenient inintegral hydrofining-hydrocracking systems employing catalysts similarto those herein described in both hydrofining and hydrocracking zones.Nevertheless, it is not necessary to introduce water into thehydrofining zone to take advantage of the results herein described. Forexample, in a single stage hydrofininghydrocracking system, the water orwater precursor can be injected alone or in combination with recyclehydrogen or hydrocracker recycle oil at an intermediate point in thereaction zone. However, procedures employing this approach willgenerally involve the use of two or more reactors in one or more stages.The distinction between stages is conventionally characterized by theseparation, or partial separation, of intermediate products such asammonia and hydrogen sulfide from the first stage product prior tointroduction of the hydrocarbon phase to the second stage. As aconsequence, these systems can involve the use of a single reaction zonein which conditions are controlled such that hydrofining occurs in thefirst portion of the catalyst bed while hydrocracking is effected in theterminal portion of the catalyst bed. In the alternative, a single stageprocess can be employed using two or more reactors in which the upstreamreactor are controlled so as to effect primarily hydrofining while thedownstream reactors are operated to promote hydrocracking. In these morecomplex plural reactor systems the advantages of this invention can berealized by injection of water or suitable water precursors to one ormore of the hydrocracking stages whether or not some hydrofining occursin each respective stage.

Particular advantage can be obtained when employing multizonehydrocracking systems wherein the organic nitrogen content of thehydrocarbon phase is gradually reduced upon sequential passage throughthe several hydrocracking zones. As the nitrogen content of thehydrocarbon phase is reduced to a level below 5 p.p.m., the beneficialeffects realized by water addition are mitigated. In fact, at nitrogenlevels substantially below 5 p.p.m., particularly 1 p.p.m. or less, theaddition of water is detrimental rather than beneficial due to itsinherent catalyst inhibiting characteristics. Consequently, it ispresently preferred to maintain the described water levels in thehydrocracking zones in which the nitrogen level in the feed issubstantially above 5 p.p.m. while avoiding water addition and/orsubstantial carryover to the terminal hydrocracking zones or stages orotherwise reducing the water content of the hydrocarbon phase thereinwhen the organic nitrogen content falls substantially below 5 p.p.m.nitrogen. By this procedure, the advantages of this invention can bemost efliciently realized while preventing any deleterious effectsattributable to the presence of water in the absence of substantialnitrogen concentrations.

As pointed out in the examples hereinafter detailed, the activity of thecatalyst immediately subsequent to addition of water or water precursorgradually approaches an increased equilibrium level and remains at thatlevel in the absence of significant catalyst deactivation as long as thewater level is maintained at the same concentration. However, when Wateraddition is discontinued after the catalyst has reached this equilibriumactivity, the hydrocarbon conversion activity of the system increasesstill further reaching a maximum and then gradually declines to theoriginal activity of the system prior to water addition. In view ofthese observations, it is possible to control the catalyst activity inseveral ways. For example, the concentration of water in thehydrocracking zone can be maintained at a predetermined constant levelthroughout the run duration. However, in that mode of operation, theactivity of the catalyst will reach an equilibrium intermediate leveland continue at that level subject only to gradual catalyst deactivationusually observed in such systerns. An alternative procedure thatprovides substantially higher average catalyst activities involvesperiodic water addition. By this procedure, a predetermined waterconcentration is maintained until the catalyst approaches equilibrium.Water addition is then discontinued to increase the catalyst activityeven further. Effective utilization of this procedure must take intoaccount the amount of water added, the time required for the catalyst toapproach its equilibrium activity and the cycle time of the catalystactivity increase and decrease subsequent to the discontinuance of wateraddition. The duration of the several cycles involved in this transientbehavior can be readily ascertained by simply adding water to theconversion zone in the prescribed amounts and observing either theincreased conversion realized at otherwise identical operatingconditions or by reducing reaction temperature sufficiently to maintaina constant conversion level until the catalyst activity has equilibratedat its higher intermediate value. The time required for the completionof this response will vary somewhat with water concentration and withthe prevailing liquid hourly space velocity. Nevertheless, the cycletime for the first step of the procedure can be readily ascertained forany given set of operating conditions by the procedure described.

After the catalyst has equilibrated at the intermediate activity level,water addition can be discontinued and the increase in catalyst activitycan be observed in a manner similar to that above described until theactivity preceding Water addition is reestablished. The data obtained bythis procedure for any system will enable the most effective utilizationof the observed phenomenon by directing the control of intermittentwater addition depending upon the cycle time of the selected system. Forexample, it may be desirable to allow the catalyst activity to approachits original value subsequent to the discontinuance of water additionbefore reintroducing water to the hydrocracking zone. However, it is notnecessary to defer the reintroduction of Water to that point. On thecontrary, additional advantage can be obtained by reintroducing water tothe system when the declining catalyst activity approaches theintermediate equilibrium value established during continuous wateraddition. By this procedure, the activity of the catalyst can bemaintained at its equilibrium level or above by intermittently addingwater or a selected Water precursor.

As previously observed, the cycle time required to traverse the completecycle of this transient catalyst activity response depends on thecharacteristics of the catalyst employed, the hydrocarbon feedstock, thenitrogen content in the hydrocracking zone, the water level employed andthe liquid hourly space velocity. However, as a general rule, whenoperating within the conditions above referred to, the time required tocomplete a cycle of this transient response is within a range of about30 minutes to about 100 hours. This range is of course narrowedconsiderably for narrower ranges of operating conditions. For example,if the preferred liquid hourly space velocity is between about 0.3 toabout 5, and the maximum water level is within a range of about 0.2 toabout 2 weight-percent, the observed cycle time will usually be withinthe range of about 1 to about 60 hours. Consequently, it is presentlypreferred to continue Water addition to the hydrocracking zone until theconversion level is stabilized which usually necessitates continuance ofwater addition for at least about 10 minutes, preferably from about 1 toabout 100 hours. The supply of added water to the reactor is thenpreferably discontinued for a period sufficient to enable the activityof the catalyst to pass through its maximum and decline to approximatelyits equilibrium value which involves periods of at least about 2 hours,usually about 1 to about 100 hours.

The-effectiveness of this procedure is demonstrated by the followingexample which is intended only to be illustrative of the concept of thisinvention and not limiting thereof.

EXAMPLE 1 The hydrocarbon feed employed in this example had thefollowing characteristics.

TABLE 1 Boiling range, by D-1160, F 555-1016 Composition, by Univ. HiMass, weight-percent:

This feed was partially hydrofined by contacting with a catalystcontaining 3 weight-percent NiO, 18 weightpercent M 3 percent P, 3percent Si-O and 73 percent alumina. Operating conditions in thehydrofining zone included a reactor temperature of 740 F., a liquidhourly space velocity of 1.0, a total reactor pressure of 2500 p.s.i.g.and a hydrogen feed rate of 8000 s.c.f./bbl. feed. The product was thenflashed to remove ammonia and hydrogen sulfide and passed to thehydrocracking zone. The described hydrofining procedures were sufficientto reduce the organic nitrogen content of the reactor charge to thehydrocracking zone to 50 p.p.m. and the organic sulfur content to 40p.p.m.

Hydrocracking conditions included a total reactor pressure of 2500p.s.i.g., a liquid hourly space velocity of 0.5 LHSV and a dry hydrogenfeed rate of 8000 s.c.f./bbl. at an overall 40 percent conversion perpass. The second zone catalyst was identical to that employed in theabove described hydrofining zone. Water was provided in the reactionzone by the addition of n-butanol to the hydrocarbon feed in amountsequivalent to 1.0 weight-percent water based on hydrocarbon.

Prior to water addition, the reaction temperature was maintained at 740for a total of 11 days. This temperature was suflicient to provide aconstant conversion per pass of 40 percent to products boiling below theinitial boiling point of the feed. Water injection was then commencedand continued for a period of 168 hours. Immediately after wateraddition the catalyst activity increased as evidenced by a drop intemperature requirement needed to maintain the desired conversion perpass. This gradual increase in catalyst activity continued for 24 hoursand then equilibrated at a level of 40 percent conversion per pass at736 F. The introduction of water was discontinued after 168 hours. Thecatalyst activity again began to increase and continued to increase fora period of 12 hours at which time the maximum catalyst activity wasobtained. The activity of the catalyst at this point was sufficient toobtain the same conversion per pass observed prior to water introductionat a reaction temperature of 710 F., 34 F. below the original reactiontemperature. However, after this point, the activity began to declineand continued its decline gradually for about 24 hours until the systemequilibrated at a conversion level of 40 percent conversion per pass at744 F. These conditions were identical to those observed before wateraddition. No significant variation in product quality or distributionwas observed at any time throughout the cycle.

From these observations, it is readily apparent that substantialadvantage can be achieved by the described procedures. These resultswere apparently the result of the influence of water in mitigating theinhibiting effect of nitrogen on both hydrogenation and hydrocrackingactivity. The aromatics hydrogenation activity of the illustrated systemincreased along with the hydrocracking activity as evidenced by theproportionately higher aromatics conversion rate which is 'known tonecessitate hydrogenation.

EXAMPLE 2 The catalyst employed in this example comprised a refractoryoxide base consisting of 20 weight-percent stabilized zeolite Y, 65weight-percent Harshaw alumina and 15 weight-percent peptized catapalalumina. The active metal components were nickel and tungsten present inamounts equivalent to 4.1 weight-percent NiO and 22.8 weight-percent W0based on the total dry weight of the active metal components andrefractory oxide. The resulting composition had an apparent bulk densityof 0.957 gram/cc. and a surface area of 248 meters per gram. Thecatalyst was sulfide by contacting with a 10% mixture of hydrogensulfide in hydrogen suflicient to convert substantially all of the metaloxides to the corresponding sulfides.

The, feed employed in this operation had the following characteristics.

TABLE 2 Boiling point range, by D-1160, F. 5 -986 Composition, by Univ.High Pass, wt.-percent:

Total saturates 45.1 Total aromatics 27.7 Sulfur, X-ray, wt.-percent2.91 Nitrogen, total, wt.-percent 0.082 AP I gravity, D-287 22.3

Conversion was effected in la recycle system in which theabove-described feed was passed into admixture with a recycle stream andthe combination was then introduced into a hydrocracking zone in whichthe described catalyst was retained in a fixed bed. The product from thereaction zone was admixed with water (during those periods in whichwater injection was employed) and the combination was flashed to producea vapor phase comprising primarily hydrogen, water vapor, andhydrocarbons having 3 carbon atoms per molecule or less. This vaporphase was employed as hydrogen recycle to the hydrocracking zone withhydrogen makeup being added as required to maintain the desired hydrogenlevels in the reactor. The water concentration in the hydrogen recyclestream Was determined by the saturation level of water in the vaporphase and was equivalent approximately 0.1 weight-percent water based ontotal hydrocarbon feed to the hydrocracking zone.

The liquid hydrocarbon phase from the separation stage was treated bycaustic scrubbing and fractionation to recover an overhead producthaving an end boiling point of 545 F. ASTM and a recycle product boilingabove 545 which was continuously recycled to the hydrocracking zone.

The hydrocracking stage was continuously operated at a pressure of 1990p.s.i.g., a liquid hourly space velocity of 1.5 and a hydrogenconcentration equivalent to 7000 standard cubic feet of hydrogen perbarrel of total feed. Reactor temperature was varied as required tomaintain a 50% conversion level per pass to products boiling below 545F. By this means the effect of water in the reaction zone was determinedas a function of the temperature required to maintain the establishedconversion level of 50% per pass.

The run was initiated and continued for a period of 9 days (216 hours)with water injection at a rate equivalent to 0.1 weight-percent based ontotal reactor charge. Water injection via saturation of the hydrogenrecycle stream was then discontinued by omitting water introduction intothe reactor product as previously described. Within 3 hours followingthe discontinuance of water injection the equivalent catalyst activityhad increased to a point at which 50% conversion per pass could bemaintained at a temperature of 714 After reaching this maximum activitythe catalyst being to deactivate and continued deactivation graduallyfor a period of about 33 hours after which time a temperature of 733 F.was required to maintain the same conversion level. Water injection wascommenced after ll days into the run (252 hours) after which time thecatalyst activity began to gradually increase and reached an activityafter an additional 48 hours sufficient to maintain 50% conversion perpass at a temperature of 722 F. The total system lined out at thattemperature (722 F.) indicating a gradual catalyst deactivation duringthe 4 day test period equivalent to a daily temperature increaserequirement (TIR) of 0.714 F. Obviously the gradual catalystdeactivation indicated by the noted temperature increase requirementtended to normalize the effect of water addition and deletion.Consequently, it becomes apparent that the differences in reactiontemperature required to maintain 50% conversion would have been evengreater than those determined in this operation were it not for thegradual catalyst deactivation observed. Nevertheless, the unusual cycleof catalyst activity resulting from the removal of water from ahydrocracking system was apparent from these observations. The catalystactivity first increased by an amount equivalent to a reactiontemperature difference of 5 F. and then gradually decreased by an amountequivalent to about 19 F. Activity then gradually increased to thepre-run activity after reintroduction of 0.1 weight-percent water(taking TIR into account). These results are summarized in Table 3.

TABLE 3 Water level,

percent Run length, hrs.

B Required to maintain 50% conversion per pass to 545 F. minus prodnet;in view of catalyst deactivation equivalent to adaily TIR at 0.7 F.

If desired, reactor temperature can be varied while cycling wateraddition and deletion to maintain a constant conversion level. Themarkedly reduced reaction temperatures required to accomplish thisresult are known to greatly reduce catalyst deactivation rate andconsequently extend the run length obtainable before regeneration isnecessitated. Conversely, reaction conditions can be maintained constantwhile cycling water addition as described to vary the overall conversionlevel. Depending upon the downstream systems necessary to post-treat thehydrocracker eflluent one or the other of these procedures may bepreferred. In most situations it is desirable to maintain a constantconversion level in view of the complexities involving in modifying theoperation of downstream fractionation and recycle systems. Consequently,in most situations it may be simpler to vary reaction temperature as afunction of water addition and conversion level in order to maintain aconstant conversion level at all times. In any event, considerableimprovement in hydrocarbon conversion can be achieved by either of theseprocedures or other alternatives readily apparent to one skilled in theart in view of the aforegoing disclosure in the appended claims.

We claim:

1. The method of hydrocracking hydrocarbons containing at least about 10p.p.m. organically bound nitrogen including the steps of contacting saidhydrocarbons at conditions of temperature, pressure and hydrogen partialpressure sufficient to effect substantial hydrocracking of saidhydrocarbons with a hydrogenation active catalyst comprising acatalytically active amount of at least one active component selectedfrom the metals, oxides and sulfides of Groups II, IV-B, VI-B, VII-B andVIII and a foraminous refractory oxide in the presence of a controlledamount of water and/or water precursor corresponding to at least about0.1 weight-percent water based on said hydrocarbon.

2. The method of claim 1 wherein said hydrocarbon contains about 10 toabout 1000 p.p.m. nitrogen as organonitrogen compounds and saidforaminous refractory oxide is selected from alumina, silica, zirconia,magnesia and combinations thereof.

3. The method of claim 1 wherein said hydrocarbons are hydrocracked inthe presence of said catalyst at a temperature of about 400 to about 900F. in the presence of at least about 500 standard cubic feet of hydrogenper barrel of said hydrocarbon.

4. The method of claim 1 wherein said refractory oxide comprises atleast one of alumina, silica-alumina cogels, silica-magnesia andzeolitic aluminosilicates, and said hydrocarbon boils primarily aboveabout 200 F.

5. The method of claim 1 wherein said hydrocarbon boils above about 200F. and contains about 10 to about 1000 p.p.m. nitrogen as organonitrogencompounds, said foraminous oxide is selected from alumina,silica-alumina cogels, silica-magnesia and zeolitic aluminosilicates,said active component comprises at least one of the metals, oxides andsulfides of calcium, magnesium, molybdenum, tungsten, manganese,rhenium, nickel, cobalt, platinum and palladium and said hydrocarbon iscontacted with said catalyst at a temperature of about 400 to about 900F. with at least about 500 standard cubic feet of hydrogen per barrel ofsaid hydrocarbon in the presence of about 0.1 to about 5 weight-percentwater.

6. The method of claim 5 wherein said hydrocarbon is contacted with saidcatalyst at a total pressure of at least about 1000 =p.s.i.'g. and aliquid hourly space velocity of 0.2 to about 10.

7. The method of claim 1 wherein said foraminous oxide comprises atleast one of alumina, silica-alumina and zeolitic aluminosilicates, saidactive component comprises at least one of the metals, oxides andsulfides of molybdenum, tungsten, nickel, cobalt, platinum andpalladium, said hydrocarbon boils primarily above about 200 F. andcontains about 10 to about 600 p.p.m. nitrogen as organonitrogencompounds and is hydrocracked in the presence of said catalyst at atemperature of about 400 to about 900 F. and a liquid hourly spacevelocity of about 0.2 to about 10 with at least about 500 standard cubicfeet of hydrogen per barrel of said hydrocarbon in the presence of about0.1 to about 2 weight-percent water.

8. The method of claim 1 wherein said foraminous oxide is selected fromalumina, silica stabilized alumina and zeolitic aluminosilicates andsaid catalyst further comprises at least about 0.2 weight-percentphosphorus.

9. The method of claim 1 wherein said catalyst comprises at least one ofalumina, silica stabilized alumina and aluminosilicates containing apromoting amount of at least one of nickel and cobalt metals, oxides andsulfides and a promoting amount of at least one of molybdenum andtungsten metals, oxides and sulfides 15 and at least about 0.2weight-percent phosphorus and said hydrocarbon is contacted with saidcatalyst at a temperature of about 500 to about 900 F. in the presenceof. at least about 500 standard cubic feet of hydrogen per barrel ofsaid hydrocarbon sufiicient to hydrocrack at least a substantialproportion of said hydrocarbon.

10. The method of hydrocrac'king hydrocarbons boiling between about 200and about 1200 F. containing about to about 1000 p.p.m. nitrogen asorganonitrogen compounds which comprises contacting said hydrocarbonswith a hydrogenation active catalyst comprising a promoting amount of atleast one of nickel, cobalt, molybdenum and tungsten metals, oxides andsulfides distended on a refractory oxide support comprising at least oneof alumina, silica, magnesia and aluminosilicates in the presence of acontrolled amount of Water and/or water precursor corresponding to about0.2 to about 2 weight-percent water at a temperature of about 400 toabout 900 F., a pressure of at least about 1000 p.s.i.g. and a liquidhourly space velocity of about 0.2 to about 10 in the presence ofhydrogen added at a rate of at least about 1000 standard cubic feet perbarrel of said hydrocarbon.

11. The method of claim 10 wherein. said catalyst comprises about 0.5 toabout 10 weight-percent of at least one of nickel and cobalt metals,oxides and sulfides determined as the corresponding oxide and at about 4to about 30 weight-percent of at least one of molybdenum and tungstenmetals, oxides and sulfides determined as the corresponding oxidedistended on said refractory oxide support comprising at least about 60percent alumina.

12. The method of claim 11 wherein said catalyst further comprises atleast about 0.2 weight-percent phosphorus, said hydrocarbon boilsprimarily above about 400 and is contacted with said catalyst at atemperature of at least 500 F. and under conditions 16 of pressure andhydrogen concentration suflicient to bydrocrack a substantial proportionof said hydrocarbon.

13. The method of claim 10 wherein said hydrocarbon is contacted withsaid catalyst over a protracted run length and said water is provided tothe system for a period of at least about 10 minutes in the form of atleast one of free Water and Water precursors convertible to water underthe conditions at which said hydrocarbon is contacted with saidcatalyst.

14. The method of claim 10 in which said water is provided to saidsystem for a period of at least about 1 to about hours in the form of atleast one of free water and alcohols convertible to water and theconditions under which said hydrocarbon is contacted with said catalyst.

References Cited UNITED STATES PATENTS 3,546,100 12/1970 Yan 2081113,023,159 2/ 1962 Ciapetta et al. 208- 3,037,930 6/1962 Mason 208111 XR3,058,906 10/1962 Stine et al. 208-111 3,157,590 11/1964 Scott, Jr., etal. 208-111 3,197,397- 7/1965 Wight 208111 3,238,120 3/ 1966 Sale208-111 3,242,067 it 3/ 1966 Arey, -Jr. et al. 208-111 3,278,417 10/1966 Van Driesen 208--109 XR 3,531,396 9/ 1970 Messing et al. 208-1113,173,853 3/1965 '-P'eralta 208.112 XR 3,501,396 3/1970 Gatsis 208112 XRTOBIAS E. LEVOW, Primary Examiner P. F. SHAVER, Assistant Examiner US.Cl. X.R.

